Hydrocracking process with rejuvenated catalyst

ABSTRACT

SILICEOUS ZEOLITE CATALYST COMPRISING ZEOLITIC MONOAND/OR DIVALENT METAL CATIONS AND A NON-ZEOLITIC GROU VIII METAL HYDROGENATING COMPONENT SUPPORTED THEREON, WBICH CATALYSTS HAVE UNDERGONE DAMAGE BY THERMAL AND/ OR HYDROTHERMAL STRESSES RESULTING IN THE MALDISTRIBUTION OF THE METAL COMPONENTS, ARE REJUVENTATED IN ACTIVATIVE BY A SEQUENTIAL TREATMENT WITH AN AQUEOUS AMMONIUM SALT TO EXCNANGE OUT AT LEAST A PORTION O THE ZEOLITIC MONOAND/OR DIVALENT METAL IONS, AND WITH AQUEOUS AMMONIA TO EFFECT A REDISTRIBUTION OF THE GROUP VIII METAL. THE TREATMENTS MAY BE PERFORMED IN EITHER ORDER.

UnitedSta-tes Patent O 3,835,028 HYDROCRACKING PROCESS WITH REJUVENATEDCATALYST John W. Ward, Yorba Linda, and Danford E. Clark,

Fountain Valley, Califi, assignors to Union Oil Company of California,Los Angeles, Calif.

No Drawing. Original application Oct. 29, 1970, Ser. No. 85,241, nowPatent No. 3,692,692, dated Sept. 19, 1972. Divided and this applicationAug. 10, 1972, Ser. No. 279,712

Int. Cl. Cg 13/02; Blllj 11/02 US. Cl. 208-111 9 Claims ABSTRACT OF THEDISCLOSURE RELATED APPLICATIONS This application is a division of Ser.No. 85,241, filed Oct. 29, 1970, now US. Pat. No. 3,692,692.

BACKGROUND AND SUMMARY OF CONVENTION It is well known that maximumactivity of the Group VIII metals for hydrogenation reactions dependsupon maintaining the metal in a finely divided state such that there isa maximum ratio of surface area to mass. Perhaps the most common methodof achieving a high degree of dispersion involves impregnating salts ofthe Group VIII metals upon porous solid supports, followed by drying anddecomposing of the impregnated salt. On nonzeolitic supports, the dryingand calcining operations often bring about a substantial migration andagglomeration of the impregnated metal, with resultant reduction inactivity. In more recent years, with the advent of highly activecrystalline zeolite catalysts of the alnminosilicate type, it has becomecommon practice to ion-exchange the desired metal salt into the zeolitestructure in an attempt to achieve an initial ionic bond between eachmetal atom and an exchange site on the zeolite, thus achieving theultimate in dispersion of metal while also bonding the metal to thezeolite in such manner as to minimize migration and agglomeration duringthe drying and calcining steps, during which at least a portion of themetal is oxidized and converted to a non-zeol-itic form. This ionexchange technique is particularly desirable in the case ofdual-function catalysts such as hydrocracking catalysts wherein it isdesirable to maintain an active hydrogenating site closely adjacent toan acid cracking site. These efforts have met with varying degrees ofsuccess.

Even though the above described ion-exchange techniques can give a highdegree of initial dispersion of Group VIII metal on the support,conditions encountered during subsequent use of the catalyst may bringabout a maldistribution of the metal with resultant reduction inactivity, entirely independent of normal deactivating phenomena such ascoking, fouling, poisoning, etc. Overheating, or contact with excessivepartial pressures of water vapor at high temperatures, such as may occurduring oxidative regeneration of the catalyst or during prolongedcontacting with hydrocarbon feedstocks, may bring about migration of theactive metal away from the exchange sites, and this migration may, underparticularly severe conditions, ultimately result in macroagglomerationof the metal into crystall'ites of 200 A. or more in.

diameter. This particular type of damage is most apt to occur underoxidizing conditions at temperatures of 500- correcting non-zeoliticGroup VIII metal maldistribution resulting from thermal and/orhydrothermal stresses encountered by the catalyst in normal usage,regeneration, or during accidental upsets cut-ailing uncontrolledtemperatures and/ or water vapor partial pressures. Normally thesestresses bring about a maldistribution of active metal short ofextensive agglomeration to particle sizes larger than about 50 A. Forexample, metal atoms or aggregates initially located closely adjacent toactive exchange sites on the carrier may migrate to other less activeareas, thus reducing the statistical likelihood of conjoint action onthe feedstock molecules of both an acidic cracking site and ahydrogenation site. Further migration may tend to drive the metal deeperinto the support structure, or into pore structures which are relativelyinaccessible by feed molecules, all resulting in reduced overallhydrogenation activity.

Limited migration of these types may occur when the catalyst, in asulfided condition (as e.g., in normal use for hydrocracking), or in anoxidized state (as during regeneration), comes into contact for morethan about 30 minutes with water vapor of greater than about 10 psi.partial pressure at temperatures above about 500 F. Extended contactingunder these conditions, or at extremely high partial pressures of watervapor, e.g., above about 100 p.s.i., can ultimately lead tomacro-agglomeration of the type previously described. If this shouldoccur, the rejuvenation procedure of this invention in some cases isless etfective per se, but can in any case be advantageously utilizedfollowing partial redispersal of the agglomerated metal by, for example,the methods described in US. Pat. Nos. 3,197,399 and/0r 3,287,257. Theprocesses described in these patents, involving respectively,alternating oxidation-reduction cycles, and alternatingsulfidingoxidation cycles, can bring about a substantial redispersion ofagglomerated metal into particles of less than about 50 A. diameter, butdo not in most instances bring about a complete recovery of the freshcatalyst activity. The process of this invention is designed to achieveat least a complete recovery of fresh activity; but in nearly all casesit is found that the rejuvenated catalysts actually exhibit greater thanfresh activity.

In the case of catalysts which originally contained a diflicultlyreducible zeolitic monovalent and/or divalent metal such as sodium,calcium, magnesium, nickel, manganese or the like, it has been foundthat the above described conditions encountered during use of thecatalyst also appear to bring about, in addition to migration of thenon-zeolitic hydrogenating metal, a detrimental redistribution of thezeolitic metal cations. Residual zeolitic metal cations, particularlysodium, are believed to occupy mainly the relatively unavailableexchange sites in the hexagonal prisms and sodalite cages of theoriginal zeolite structure, but under the described conditions of use,migration to more active cracking sites appears to Occur with resultantloss in cracking activity. Divalent metal cations such as the alkalineearth metals, which may have been originally exchanged into the zeoliteto achieve hydrothermal stability, may also migrate to undesirablesites, and in any event appear after extended use of the catalysts underthe described conditions to be no longer necessary for stabilization,the anionic zeolite structure having acquired stability by virtue of thethermal and/0r Patented Sept. 10, 1974 hydrothermal conditions whichbrought about the cation migration. It is hence desirable in the case ofthese damaged catalysts to remove zeolitic monoand/or divalent metalcations, in addition to redistributing the non-zeolitlc Group VIII metalhydrogenating component. The former is accomplished by ammonium ionexchange; the latter by the ammoniation step.

As employed herein, the term non-zeolitic metal refers to the metalcontent of the catalyst, other than anionic lattice metals such asaluminum, which is not chemically bonded to the anionic exchange sitesof the zeolite, while conversely, zeolitic metal refers to the metalcontent which is so bonded. The easily reducible metals such as theGroup VIII noble metals are normally present primarily as non-zeoliticmetal, as a result of previous reduction with hydrogen, oxidation and/or sulfiding treatments. The difficultly reducible metals such as thealkali and alkaline earth metals are normally present almost exclusivelyas zeolitic cations, since they are not affected by the normalreduction, oxidation or sulfiding treatments. Metals of intermediatereducibility such as nickel, copper and the like may be present in bothzeolitic and non-zeolitic form.

Briefly stated, the complete rejuvenation procedure of this inventioninvolves two essential steps performed sequentially in either order:

(1) An aqueous ion exchange treatment with an ammonium salt solution toeffect replacement of at least a portion of the detrimental zeoliticmonoand/or divalent metal cations with ammonium ions, thereby increasingthe cracking activity; and

(2) A treatment with ammonia and water to redistribute the non-zeoliticGroup VIII metal, thereby increasing the hydrogenation activity.

The sequence of ammoniation followed by ion exchange is preferredbecause a more active catalyst is usually obtained than by the oppositesequence.

While we do not wish to be bound by any theoretical explanation of theresults achieved herein, it appears that the ammoniation step involves areconstitution of the original ion-exchange cationic species of thehydrogenating metal, or the soluble metal ammino-hydroxide, in the poresof the catalyst by treatment with aqueous ammonia. Hydration andammoniation of the deactivated catalyst, in which the Group VIII metalis in an oxidized or sulfided form, fills the micropores with a strongaqueous ammonia solution. This results in dissolution of the Group VIIImetal oxide or sulfide in the ammonia solution to form the originalcationic species, or soluble ammino-hydroxide, which was originallyion-exchanged into the zeolite. For example, palladium oxide on thezeolite support will form the Pd(NH ion, which then migrates back to theoriginal ion exchange sites. The original distribution of palladium isthen theoretically obtained after drying and recalcining as in theoriginal catalyst preparation method. Similarly, platinum oxide onamorphous silica-alumina will form Pt(NH )4(OH) or Pt(NH (OH) which,being stronger bases than NH OH, will tend to combine with the originalacid sites on the support. The original distribution of platinum withrespect to acid sites will then theoretically be obtained after dryingand calcining.

Irrespective of the correct theoretical explanation for the resultsachieved herein, the experimental evidence available indicates that theprocedures described herein can give complete rejuvenation ofzeolite-based Group VIII metal catalysts wherein a maldistribution ofmetals has occurred as a result of overheating, or of contacting thecatalyst while in an oxidized or sulfided state with water vapor attemperatures between about 500 and 1200" F.

4v 7 DETAILED DESCRIPTION (A) Hydration-Ammoniation This portion of therejuvenation may be carried out by any desired procedure which will givea substantial adsorption of water into the micro-pores of the catalystand absorption-solution of a substantial proportion, at least about 5weight-percent and preferably 10 to 35 .weightpercent, of ammonia intothe adsorbedwater phase, based on the weight of the water phase.Preferably, the catalyst is first hydrated and then ammoniated, butsimultaneous hydration and ammoniation is also contemplated.Simultaneous hydration and ammoniation can be effected by wetting thecatalyst with an aqueous ammonia solution. In the preferred procedure,the catalyst is simply hydrated in moist air to an extent of e.g. 5-40weight-percent, and then contacted with gaseous ammonia until the liquidWater phase is substantially saturated with ammonia. The hydration andammoniation steps are pref erably carried out at temperatures between 0and F., but temperatures up to about 300 F. or even higher arecontemplated. Normally these steps are carried out at atmosphericpressure, but reduced or superatmospheric pressures may be utilized.

For treating large batches of catalyst, it is normally desirable tohydrate by passing moist air or other Wet gas through a bed of thecatalyst until there is a substantial breakthrough of water vapor in theefliuent gases. Ammoniation may be similarly effected by passingammonia, or ammonia-containing gases through the bed until ammoniaappears in the off gases. Simultaneous hydration and ammoniation can beeffected by passing a gas stream containing both ammonia and water vaporthrough the catalyst bed until both water vapor and ammonia appear inthe ofif gases.

In any of the above procedures, it will be understood that in caseswhere the zeolite base is in a hydrogen or decationized form, theammoniation will at least partially convert the zeolite to an ammoniumzeolite, in addition to saturating the adsorbed water with ammonia. The

dissolved ammonia and the zeolitic ammonium ions are.

hence the drying and calcining steps can be commeced y substantiallyimmediately thereafter. However, in some cases, as for example where asubstantial agglomeration of metal has occurred, it may be desirable toage the. catalyst in its hydrated-ammoniated form for periods rangingfrom about one hour to twelve hours or more.

Regardless of whether the ammoniation step is performed before or afterthe ammonium ion exchange step, it is ordinarily necessary to convertthe hydrated-ammoniated catalyst to a dehydrated, deammoniated, oxidizedform. These objectives can be achieved with difficulty by a carefullycontrolled rapid heatup to, e .g., 950 F. in air, but to achieve maximumcatalytic activity in this manner would be a practical impossibility.The reason for this stems from the observed fact that at temperaturesbetween about 500 and 950 F. the Group VIII metal on the catalyst, whenin an oxidized state, tends to undergo severe agglomeration unless thewater vapor partial pressure is carefully controlled. Hence, a rapidheatup in air would tend to raise the catalyst temperature to above 500F. before some portions of the catalyst bed (or even some areas of eachcatalyst pellet) had been sufficiently dehydrated to permit control oflocalized water vapor concentrations. In general, in order to avoidagglomeration of oxidized metal on the catalyst in the 500- 950 F.temperature range, itis desirable to maintain Water vapor partialpressures below about 10 psi, and preferably below 2 p.s.i. It istherefore highly desirable to reduce the water content of the catalystto a practical minimum at temperatures below 500 F., for at temperaturesabove about 500 F. the catalyst is rapidly being converted to anoxidized state with chemical evolutioncf water. At below about 500 F.,the metal or metal oxide is not affected by water vapor.

Accordingly, for the above purposes, a preferred drying step is carriedout by passing a stream of air or other non-reducing gas through a bedof the catalyst without maintaining dewpoint control over the efliuentgases. It is generally preferable to start the drying at a lowtemperature of e.g., 100 to 200 'F., and incrementally raise thestripping gas temperature to a level in the 300 F. to 500 F. range.During the drying step, nearly all of the ammonia absorbed into thewater phase in the catalyst is removed, any remaining ammonia beingprimarily in the form of zeolitic ammonium cations. It is this zeoliticammonium which creates an additional problem of water vapor partialpressure control during the subsequent calcination step, for it isduring this step that zeolitic ammonium is oxidized to form additionalwater vapor (and nitrogen), which adds its effect to that of the watervapor generated by desorption of any remaining water in the catalyst.Hence the desirability of stripping out at least about one-half, andpreferably at least about two-thirds, of the adsorbed water during thedrying step at temperatures below about 500 F.

It is to be noted also that reducing gases such as hydrogen should besubstantially absent during the drying step. For reasons which are notclearly understood, direct reduction of the complex metal ammine cationto the free metal always results in severe agglomeration thereof. Hencethenecessity for first converting the metal ammino complex to anoxidized state during the calcining step, and then later reducing theoxidized metal to activate the same for use in hydrocarbon conversions.Suitable stripping gases for use in the drying step include air or otheroxygen-containing gases, nitrogen, argon, methane and the like. Thedrying is normally carried out at atmospheric pressures, or slightlyelevated pressures of e.g., 50 to 100 p.s.i.g. Where crystalline zeolitecatalysts are concerned, it is normally desirable to reduce the watercontent to about 5-10 weight-percent.

The calcination step may be performed in the same apparatus employed forthe drying step if desired, e.g., in a rotary kiln, a moving beltfurnace, or in a vessel containing a fixed bed of the catalyst. Toinitiate the calcination, air is admixed with the stripping gas,initially.

in small proportions to provide an oxygen concentration of e.g., about0.1% to 1% by volume. The temperature of the calcination gas is thengradually increased from about 500 F. to 700750 F. while graduallyincreasing the oxygen concentration to e.g., about .5 to 2%. During theentire heatup period, water concentration in the calcination vesselshould be carefully controlled, as by monitoring the efiiuent gases tomaintain a dew-point below about 40 F., preferably below 20 F. Followingeach incremental increase in oxygen concentration it is generallydesirable, in the case of fixed bed calcinations, to wait for theexothermic temperature wave to pass through the catalyst bed and untiloxygen breakthrough has occurred before the next incremental increase inoxygen concentration is effected. Continuing in this manner, inlet gastemperatures and oxygen concentrations are increased until temperaturesof about 900 to 1100 F. and final oxygen concentrations in the range ofabout 2-10% or more are reached. When the terminal temperature andoxygen concentrations are reached, the calcination is then preferablycontinued for a sufiicient length of time to give an effluent gas streamhaving a dewpoint below about 0 F., preferably below about -20 F.

The gradual heatup procedure With incremental in creases in oxygenconcentration as described above is a practical necessity when thecalcination is carried out with a fixed bed of catalyst through whichthe calcination gases are passed. It is not intended however that theinvention be limited to this procedure, for a considerably more rapidheatup at high oxygen concentrations can be utilized when the catalystis arranged in thin layers through which the gases pass, whereby theeffect of water vapor on downstream portions of the catalyst isminimized. Commercially, a rotary kiln equipped with litters and a dryair sparger to provide good ventilation of the cascading bed of catalystis very effective in achieving the desired results of this invention. Aparticularly critical period during the calcination appears to be theperiod of burnoff of zeolitic ammonium ions, which occurs primarily attemperatures above about 750 F, and can generate a burning wave in thecatalyst wherein instantaneous temperatures and water vaporconcentrations may inhibit full recovery of the original fresh catalystactivity. Accordingly, greatest care should be exercised to minimizewater vapor concentrations during the 7501000 F. heating cycle:

(B) Ammonium Ion Exchange Ion exchange with ammonium ions may be carriedout by conventional procedures which involve in general contacting thecatalyst with an aqueous solution of any desired ammonium salt, e.g.,the nitrate, sulfate, chloride, acetate, or the like. Preferablyammonium nitrate is employed, and preferred salt concentrations rangebetween about 5% and 50% by weight. Practical contacting temperaturesrange between about 10 and 100 C., preferably 20-90". The lowertemperatures of about 20-40" C. are preferred from the standpoint ofminimizing the leaching out of Group VIII metal from the catalyst,although higher temperatures give more rapid exchange. The ion exchangeefiiciency is further enhanced by acidifying the exchange solution to apH of about 3-5.5 with an acid, e.g., acetic acid.

The contacting may be carried out in a single stage, in plural batchstages, or continuously by flowing a stream of the ammonium saltsolution through a bed of the catalyst. Normally it is desirable tocontrol the severity, or use the number of stages required, to remove atleast about 50%, preferably at least about of the total zeolitic monoanddivalent metal content.

Following the ion exchange step, the catalyst is washed to removeresidual salts, and, if the ammoniation step was carried out priorthereto, is subjected to final drying and calcining as previouslydescribed. However, if the exchange step is performed first, thecatalyst need only be dried to the desired water content prior toammoniation, although in some cases better overall activity is obtainedif an intervening calcination is performed.

(C) Catalyst Compositions Catalyst compositions which may be rejuvenatedby the above procedures include hydrogenation catalysts, hydrocrackingcatalysts, isomerization catalysts, reforming catalysts and the likewhich comprise a Group VIII metal, with or without other metals or metaloxides such as those of the Group VIB metals, supported on a siliceouszeolite base having an ion exchange capacity of at least about 0.01meq./gm., and preferably at least about 0.1 meq./ gm. Suitable siliceouszeolite bases include for example the crystalline aluminosilicatemolecular sieves such as the Y, X, A, L, T, Q, and B crystal types, aswell as zeolites found in nature such as for example mordenite,stilbite, heulandite, ferrierite, dachiardite, chabazite, erionite,otfretite, and the like. The preferred crystalline zeolites are thosehaving crystal pore diameters between about 7l5 A., wherein the SiO /AlO mole ratio is about 3/1 to 10/1. For most catalytic purposes, e.g.,catalytic hydro cracking, it is preferable to replace most or all of thezeolitic sodium normally associated with such zeolites with othercations, particularly hydrogen ions and/or polyvalent metal ions such asmagnesium, calcium, zinc, rare earth metals and the like.

The utilitarian effect of the ammonium ion exchange treatment of thisinvention is most evident in the case of catalysts containingsignificant proportions, e.g., 0.5-% by weight, of zeolitic mono and/ordivalent metal ions, particularly the metals of Groups 1A, HA and11B,-e.g.. sodium, potassium, calcium, magnesium, zinc, 'etc., as'wellas iron, cobalt, nickel and the like.

In addition to the crystalline zeolite bases describe above, otherZeolitic bases may be employed such as the zeolitic cogels of silca andalumina, silica and titania, silica and zirconia, silica and magnesiaand the like.

The Group VIII metal hydrogenating component is ordinarily added to thezeolite base by ion exchange with an aqueous solution of a suitablecompound of the desired metal wherein the metal is present in a cationicform, as described for example in US. Pat. No. 3.236.762. Suitableamounts of the iron group metals, i.e., iron, cobalt and nickel, mayrange between about 1% and by weight, while the noble metals, e.g.,palladium and platinum are normally employed in amounts ranging betweenabout 0.1% and 2% by weight. The noble metals, particularly palladiumand platinum, are normally preferred herein. Other metals such asrhenium may also be included.

When catalysts of the foregoing description are utilized for extendedperiods of time at temperatures of, e.g., 400950 F. in hydrocarbonconversions such as hydrocracking, hydrogenation, isomerization,reforming and the like, a progressive decline in catalyst activitynormally occurs as a result of coke deposition. A more rapid or suddendecline in activity will normally follow when the catalyst encounters,either during hydrocarbon conversion or during regeneration, any of theadverse conditions of temperature and water vapor partial pressurepreviously described. Deactivation by coking is normally almostcompletely reversible by conventional oxidative regeneration attemperatures of e.g., 7-50-1100 F. When it is found that such oxidatixeregeneration restores less than about 90% of the fresh hydrogenationactivity, and less than about 90% of the fresh cracking activity, it maybe assumed that some undesirable maldistribution of the metal contenthas occurred, such as to warrant use of the rejuvenation proceduresdescribed herein. It will be understood that hydrogenation activity ismeasured in terms of, and is inversely proportional to, the volume ofcatalyst required to effect a given degree of hydrogenation per pass ofa particular compound, e.g., benzene, at a particular set ofhydrogenation conditions. Cracking activity can be measured in terms ofthe standard Cat-A cracking activity index.

The following examples are cited to illustrate the invention, but arenot to be construed as limiting in scope.

EXAMPLE I This example illustrates a typical type of hydrothermaldeactivation which can occur during catalytic hydrocracking. Ahydrocracking run was carried out over a period of about twenty monthsutilizing a catalyst consisting of 0.5 weight-percent Pd supported on aY molecular sieve cracking base having a SiO /Al O mole-ratio of about4.7, wherein about of the zeolitic ion exchange capacity was satisfiedby magnesium ions (3 weight-percent MgO), about 10% by sodium ions, andthe remainder (55%) by hydrogen ions. This catalyst was maintained in asulfided condition throughout the run by virtue of a sour recycle gascontaining about 0.3 volume-percent of hydrogen sulfide. The run wascarried out at a pressure of about 1500 p.s.i.g., with space velocitiesvarying between about 1.3 and 1.7, hydrogen rates varying between 5,000and 7,000 s.c.f./b., and with hydrocracking temperatures progressivelyincreasing from about 500 F. to 680 F. The feedstock was a substantiallysulfurand nitrogenfree unconverted gas oil (400-850" F. boiling range)derived from a previous stage of hydrocracking. Hydrocrackingtemperatures were incrementally raised during the runto maintain 6070volume percent conversion per pass to gasoline.

. During this run, a foaming problem was encountered in the recycle gaswater-washing column. resulting in a substantial quantity of water beingcarried into the reactor, giving an estimated 100 p.s.i. partialpressure of water vapor therein for a period of about 4 hours. Animmediate temperature increase of about 5.5 F. was required in order tomaintain the desired conversion level, this temperature increasecorresponding to a loss in catalytic activity of about At the end ofthis run, the catalyst was carefully regenerated by oxidative combustionat temperatures ranging from about 700 up to 1000 F., utilizing aregeneration gas comprising oxygen in amounts increasing from about 0.1to 3.0 volume percent, whereby water vapor partial pressures weremaintained at a value below about 0.25 p.s.i.a. at all regenerationtemperatures above 500 F. The regenerated catalyst was then tested foractivity compared to that of the fresh catalyst. The feedstock used forthe activity test was the same feed used in the previous hydrocrackingrun, doped with thiophene to a level of 0.48% sulfur to provide an Hs-containing atmosphere for the hydrocracking. Conditions of theactivity test were: pressure 1450 p.s.i.g., LHSV 1.7, hydrogen/oil ratio8,000 s.c.f./b., conversion per pass 52-54 volume-percent to gasoline.The following table shows the temperatures required to maintain theabove conversion as a function of time:

TABLE 1 EXAMPLE H A sample of the catalyst regenerated as describedabove was subjected to the hydration-ammoniation step of the presentinvention as follows:

(1) Allowed to hydrate in ambient air to a saturation value of about 25%by weight of water on a hydrated basis.

(2) Treated with gaseous ammonia at ambient temperatures and pressuresto substantially saturate the water in the catalyst pores (about 2530weightpercent NH based on water).

(3) Allowed to stand overnight in ambient air to volatilize most of theexcess ammonia.

(4) Stripped and partially dried to a water content of about 6-8weight-percent in a mufiie furnace through which a stream of dry air waspassed for two hours at temperatures increasing from ambient to 480 F.,and then for two hours at 480 F.

to 400 F. E.P. gasoline, F. 20 541 40 550 60 555 80 v 558 100 560 Theforegoing data shows that the ammonia-rejuvenated catalyst of thisexample was almost as active as the original fresh catalyst; therequired conversion temperatures leveled out at 100250 hours to a valueabout 9-10 higher than the corresponding temperatures required for thefresh catalyst.

. If the drying step (4) in the above example is omitted, and thecatalyst simply calcined for one hour at temperatures from ambient to930 followed by one hour at 930 E, the resulting catalyst is less activethan the original regenerated catalyst of Example I, due to palladiumagglomeration brought about by excessive water vapor partial pressuresduring the rapid heatup from ambient to 930 F.

EXAMPLE III Another sample of the regenerated catalyst from Example Iwas subjected to the ammonium ion-exchange step of the present inventionas follows:

(1) Slurried 100 g. of catalyst in one liter of 10% NH NO solution,added acetic acid to give pH=3.5, stirred at 180 F. for two hours,filtered and washed.

(2) Step (1) repeated for a total of three exhanges.

(3) Final product dried and calcined as in step (5) of Example II.Analysis showed that the ion exchange had reduced the sodium content to0.5% Na O and the magnesium content to about 0.7% MgO.

The above data shows that the ion-exchange technique alone provides a.catalyst which is initially more active than the fresh catalyst ofExample I or the ammoniated catalyst of Example II, but the activityrapidly declined so that after 150 hours it was substantially less.active than the fresh catalyst or the ammoniated catalyst. It is henceapparent that neither treatment alone fully restored the fresh catalystactivity.

EXAMPLE IV Another sample of the regenerated catalyst from Example I wasfirst subjected to the ammonium ion-exchange treatment described inExample III'and then to ammoniation as described in Example II. The sameactivity test gave the following results:

TABLE 4 Temp. for 52-54% conversion Hours: to 400 F. El. gasoline, F.516

TABLE 5 Temp. for 5254% conversion Hours: to 400 F. E.P. gasoline, F. 50504 75 509 100 511 125 513 150 514 It is thus apparent that theammoniation-ion exchange sequence gives a catalyst which is even moreactive, and which deactivates at a lower rate, than the catalystrejuvenated by the reverse order of treatment.

For convenience, the essential data from the foregoing Examples istabulated as follows:

TABLE 6 Catalyst treatment- Temp. F.) for 5254% conversion to 400 F. El.gasoline Example I I II III IV V Regen- NH4NO1, Nils-H2O, Fresh eratedNIB-H2O NH4N0; NHz-HgO NH4NO3 Activity testing of this catalyst asdescribed in Example I gave the following results:

TABLE 3 Temp. for 52-54% conversion The foregoing details as to specificcatalysts and rejuvenation conditions are not intended to be limiting ineffect. The following claims and their obvious equivalents are intendedto define the true scope of the invention.

We claim:

1. A process for the hydrocracking of a hydrocarbon feedstock to producelower boiling hydrocarbons, which comprises subjecting said feedstock inadmixture with added hydrogen to elevated conditions of pressure andtemperature in contact with a rejuvenated catalyst compositioncomprising a non-zeolitic Group VIII metal hydrogenating componentsupported on a siliceous zeolite carrier, said catalyst havingpreviously been subjected to damaging thermal" 'and/ or -'hydr othermalconditions resulting in (a) am'aidistribution of said Group VIII metalhydrogena ting' component 'on' said carrier with resultant loss inhydrogenation activityj'and (b) a deleterious redistributionofzeolitic'mon'oand/or divalent metal cations originally contained thereinwith resultant loss in cracking activity, theso damaged catalyst havingthereafter been rejuvenated by a sequential treatment in either orderof: t

.(A) contacting and damaged catalyst with water or water vapor andwithammonia at below 300. F. to etfecta substantial hydration andammoniation thereof, and thereafter drying and calcining the catalyst toeffect dehydration and deammoniation thereof; and

(B) contacting the damaged catalyst with an aqueous ammonium saltsolution to effect a substantial replacement of said zeolitic monoand/ordivalent metals with ammonium ions, and thereafter drying the catalyst.

2. A process as defined in claim 1 wherein step (A) is performed priorto step (B).

3. A process as defined in claim 1 wherein step (B) is performed priorto step (A).

4. A process as defined in claim 1 wherein said zeolitic monovalentand/or divalent metals comprise an alkali metal and/or alkaline earthmetal.

5. A process as defined in claim 1 wherein said Group VIII metal is aGroup VIII noble metal.

6. A process as defined in claim 1 wherein said siliceous zeolitecarrier is a crystalline molecular sieve.

7. A process as defined in claim 1 wherein:

(a) step (A) is performed prior to step (B);

(b) saidfzeolitic monoval ninand/o'fr 1' (c) said GroupVIII metalis"pall a'diuin"and or p inum; and 1 I (d) said siliceous zeoliteCarrie'r'is a Y molecuiar sieve wherein the zeolitic catio'ns'areprimarily hydrogen ions and/or polyvalent metalionsf' I:

8. A process as defined in claim 7 wherein said "damaging thermal and/orhydrothermal conditions were encountered during a previous hydrocrackingrun employing said catalyst, and/ or during a subsequent oxidativeregeneration following such hydrocracking run. a

9. A process as defined in claim 1 wherein said damaging thermal and/orhydrothermal conditions were encountered during a previous hydrocrackingrun employing said catalyst, and/or during a subsequent oxidativeregeneration following such hydrocracking run.

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